Upgrading raw shale-derived crude oils to hydrocarbon distillate fuels

ABSTRACT

Integrated processes for upgrading crude shale-derived oils, such as those produced by oil shale retorting or by in situ extraction or combinations thereof. Processes disclosed provide for a split-flow processing scheme to upgrade whole shale oil. The split flow concepts described herein, i.e., naphtha and kerosene hydrotreating in one or more stages and gas oil hydrotreating in one or more stages, requires additional equipment as compared to the alternative approach of whole oil hydrotreating. While contrary to conventional wisdom as requiring more capital equipment to achieve the same final product specifications, the operating efficiency vis a vis on-stream time efficiency and product quality resulting from the split flow concept far exceed in value the somewhat incrementally higher capital expenditure costs.

FIELD OF THE DISCLOSURE

Embodiments disclosed herein relate generally to upgrading of wholeshale oils, such as crude oil shale-derived oils produced by oil shaleretorting or in situ extraction.

BACKGROUND

Whole shale oil contains distillate and residuum fractions with widelydifferent compositions beginning with light hydrocarbons typicallyboiling in the C4 range and extending to a wide range of higherdistillate boiling hydrocarbons and heteroatomic compounds up to andincluding compounds boiling in the 975° F.+ vacuum residuum range. Thesewide boiling ranges of hydrocarbons and heteroatomic compounds can havewidely varying reactivities in downstream catalytic or thermal upgradingprocesses. Raw shale oil may contain nitrogen containing compounds,metals such as arsenic and/or selenium and compounds of arsenic and/orselenium and other impurities such as sulfur containing compounds.Additionally, raw shale oil may also contain particulate mattercomprised of fine oil shale particles that are entrained or trapped inthe recovered whole shale oil products from the upstream retorting or insitu extraction processes. The hydrocarbons in whole shale oil mayinclude various paraffins, olefins, diolefins, and aromatics, includingheavy oil, gas oil, and asphaltenes containing multiple fused aromaticring compounds.

Whole shale oils are conventionally processed in a single hydrotreatingreactor. The reactor is operated at a single hydroprocessing severity,which, however, cannot effectively perform the required hydroprocessingreactions to upgrade the whole shale oil without encountering theprocessing penalties of severe fouling/plugging and poor selectivity,e.g., gas formation. This is a result of the different compositions andreactivities of the various fractions in the whole shale oil.

Other processes disclose use of multiple reactors to upgrade the wholeshale oil. For example, U.S. Pat. No. 4,133,745 discloses a process forprocessing shale oil in which the shale oil is fractionated into anaphtha cut, boiling below 350° F., and a gas oil cut, boiling above350° F. The naphtha cut is then hydrotreated to remove nitrogen and thegas oil cut is treated to remove impurities, such as by caustictreating. The gas oil cut is then hydrotreated to remove nitrogencompounds and fractionated to produce a second naphtha cut boiling below450° F. While this approach manages to hydroprocess all but a minorportion of the light naphtha portion, it suffers from the attendantdeficiencies of fouling, plugging, and selectivity issues.

SUMMARY OF THE CLAIMED EMBODIMENTS

In contrast to prior processes, embodiments disclosed hereinadvantageously separate whole shale oil or partially hydrotreated wholeshale oil so as to effectively perform the required hydroprocessingreactions, namely, diolefins saturation; hydrodemetallization (HDM);monoolefins saturation; hydrodenitrogenation (HDN); hydrodesulfurization(HDS); hydrodeoxygenation (HDO), without the penalties of severefouling/plugging and poor selectivity, e.g., gas formation, to producehigher-valued hydrocarbon products.

Contrary to conventional thought in the industry, flow schemes disclosedherein, having a split flow concept for hydroprocessing of differentshale oil fractions in separate hydroprocessing reactors, may beeconomically viable, even though they may require additional reactorsand associated processing equipment. However, the ability to optimizehydroprocessing conditions in each of the separate hydroprocessingreactors will allow significant improvements such as for exampleminimization of catalyst fouling rates; higher on-stream timeefficiencies; and significant improvements in hydroprocessed productqualities with the latter affording higher product revenues, thuslowering the net operating costs and mitigating the incrementally highercapital investment costs.

In one aspect, embodiments disclosed herein relate to an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or by in situ extraction or by mixtures thereof. The processmay include the following steps: (a) fractionating the whole shale oilinto a first fraction comprising naphtha, kerosene and diesel and anatmospheric bottoms fraction comprising gas oil and residuum; (b)contacting the first fraction and hydrogen in a first-stagehydroprocessing reactor containing a hydrogenation catalyst to saturatediolefins contained in the first fraction and recovering an effluentfrom the first-stage hydroprocessing reactor; (c) feeding the effluentfrom the first-stage hydroprocessing reactor of step (b) without phaseseparation to a second-stage hydroprocessing reactor operated in anupflow mode and containing catalysts to perform hydrodemetallization andsaturation of mono-olefins in the effluent from the first-stagehydroprocessing reactor and recovering an effluent from the second-stagehydroprocessing reactor; (d) feeding the effluent from the second-stagehydroprocessing reactor of step (c) without phase separation to athird-stage hydroprocessing reactor having one or more beds of catalystto perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the effluent from thesecond-stage hydroprocessing reactor and recovering an effluent from thethird-stage hydroprocessing reactor; (e) feeding the atmospheric bottomsfraction and hydrogen to a fourth-stage hydroprocessing reactor operatedin an upflow mode and containing catalysts to performhydrodemetallization of the atmospheric bottoms fraction and recoveringan effluent from the fourth-stage reactor; (f) feeding the effluent fromthe fourth-stage hydroprocessing reactor of step (e) without phaseseparation to a fifth-stage hydroprocessing reactor having one or morebeds of catalyst each containing a catalyst to perform one or more ofhydrotreating and hydrocracking of the effluent from the fourth-stagehydroprocessing reactor and recovering an effluent from the fifth-stagehydroprocessing reactor; and (g) processing the effluents from thefifth-stage hydroprocessing reactor of step (f) and the third-stagehydroprocessing reactor of step (d) in a separation train to recover twoor more hydrocarbon fractions.

In another aspect, embodiments disclosed herein relate to an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or by in situ extraction or combinations thereof. The processmay include the following steps: (a) fractionating the whole shale oilinto a first fraction comprising naphtha, kerosene and diesel and anatmospheric bottoms fraction comprising gas oil and residuum; (b)feeding the first fraction and hydrogen to a first-stage hydroprocessingreactor containing a hydrogenation catalyst to saturate diolefinscontained in the first fraction and recovering an effluent from thefirst-stage hydroprocessing reactor; (c) feeding the effluent from thefirst-stage hydroprocessing reactor of step (b) without phase separationto a second-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; (d) feeding the atmospheric bottoms fraction and hydrogen to athird-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and recovering aneffluent from the third-stage hydroprocessing reactor; (e) feeding theeffluent from the third-stage hydroprocessing reactor of step (d)without phase separation to a fourth-stage hydroprocessing reactorhaving one or more beds of catalyst each containing a catalyst toperform one or more of hydrotreating and hydrocracking of the effluentfrom the third-stage hydroprocessing reactor and recovering an effluentfrom the fourth-stage hydroprocessing reactor; (f) feeding the effluentfrom the second-stage hydroprocessing reactor of step (c) and theeffluent from the fourth-stage hydroprocessing reactor of step (e),without phase separation, to a fifth-stage hydroprocessing reactorhaving one or more beds of catalyst to perform hydrodenitrogenation,hydrodesulfurization, hydrodeoxygenation, and aromatics saturation ofthe effluents from the second- and fourth-stage hydroprocessing reactorsand recovering an effluent from the fifth-stage hydroprocessing reactor;and (g) processing the effluents from the fifth-stage hydroprocessingreactor of step f) in a separation train to recover two or morehydrocarbon fractions.

In another aspect, embodiments disclosed herein relate to an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or by in situ extraction or combinations thereof. The processmay include the following steps: (a) contacting the whole shale oil andhydrogen in a first-stage hydroprocessing reactor containinghydrogenation catalysts to saturate diolefins contained in the wholeshale oil and recovering an effluent from the first-stagehydroprocessing reactor; (b) feeding the effluent from the first-stagehydroprocessing reactor of step (a) without phase separation to asecond-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; (c) fractionating the effluent from the second-stagehydroprocessing reactor of step (b) into a partially hydrotreatedfraction comprising naphtha, kerosene and diesel and a partiallyhydrotreated bottoms fraction comprising gas oil and residuum; (d)feeding the partially hydrotreated fraction to a third-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andhydrodearomatization of the partially hydrotreated fraction andrecovering an effluent from the third-stage hydroprocessing reactor; (e)feeding the partially hydrotreated bottoms fraction to a fourth-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the partially hydrotreated bottoms fraction andrecovering an effluent from the fourth-stage hydroprocessing reactor;and (f) processing the effluents from the third-stage hydroprocessingreactor of step (d) and the fourth-stage hydroprocessing reactor of step(e) in a separation train to recover two or more hydrocarbon fractions.

In another aspect, embodiments disclosed herein relate to an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or by in situ extraction or combinations thereof. The processmay include the following steps: (a) contacting the whole shale oil andhydrogen in a first-stage hydroprocessing reactor containing catalyststo saturate diolefins contained in the whole shale oil and recovering aneffluent from the first-stage hydroprocessing reactor; (b) feeding theeffluent from the first-stage hydroprocessing reactor of step (a)without phase separation to a second-stage hydroprocessing reactoroperated in an upflow mode and containing catalysts to performhydrodemetallization and saturation of mono-olefins in the effluent fromthe first-stage hydroprocessing reactor and recovering an effluent fromthe second-stage hydroprocessing reactor; (c) fractionating the effluentfrom the second-stage hydroprocessing reactor of step (b) into apartially hydrotreated fraction comprising naphtha, kerosene and dieseland a partially hydrotreated bottoms fraction comprising gas oil andresiduum; (d) feeding the partially hydrotreated bottoms fraction to athird-stage hydroprocessing reactor having one or more beds of catalystto perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the partiallyhydrotreated bottoms fraction and recovering an effluent from thethird-stage hydroprocessing reactor; (e) mixing the partiallyhydrotreated fraction and the effluent from the third-stagehydroprocessing reactor to form a mixture; (f) feeding the mixture to afourth-stage hydroprocessing reactor having one or more beds of catalystto perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the mixture andrecovering an effluent from the fourth-stage hydroprocessing reactor;and (g) processing the effluent from the fourth-stage hydroprocessingreactor in a separation train to recover two or more hydrocarbonfractions.

In another aspect, embodiments disclosed herein relate to an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or by in situ extraction or combinations thereof. The processmay include the following steps: (a) contacting the whole shale oil andhydrogen in a first-stage hydroprocessing reactor containing catalyststo saturate diolefins contained in the whole shale oil and recovering aneffluent from the first-stage hydroprocessing reactor; (b) feeding theeffluent from the first-stage hydroprocessing reactor of step (a)without phase separation to a second-stage hydroprocessing reactoroperated in an upflow mode and containing catalysts to performhydrodemetallization and saturation of mono-olefins in the effluent fromthe first-stage hydroprocessing reactor and recovering an effluent fromthe second-stage hydroprocessing reactor; (c) feeding the effluent fromthe second-stage hydroprocessing reactor of step (b) without phaseseparation to a third-stage hydroprocessing reactor having one or morebeds of catalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the effluent from thesecond-stage hydroprocessing reactor and recovering an effluent from thethird-stage hydroprocessing reactor; (d) fractionating the effluent fromthe third-stage hydroprocessing reactor of step (c) into a partiallyhydrotreated fraction, comprising naphtha, kerosene and diesel, and apartially hydrotreated vacuum gas oil fraction; (e) feeding thepartially hydrotreated vacuum gas oil fraction to a fourth-stagehydroprocessing reactor having one or more beds of catalyst to performhydrocracking of the partially hydrotreated vacuum gas oil fraction andrecovering an effluent from the fourth-stage hydroprocessing reactor;(f) feeding the effluent from the fourth-stage hydroprocessing reactorto the fractionation step (d).

Other aspects and advantages will be apparent from the followingdescription and the appended claims,

BRIEF DESCRIPTION OF DRAWINGS

FIGS. 1-5 illustrate simplified process flow diagrams of processes forupgrading whole shale oil according to embodiments disclosed herein.

DETAILED DESCRIPTION

In one aspect, embodiments herein relate to hydrotreating whole shaleoil such as crude oil shale-derived oils produced by either oil shaleretorting or in situ extraction or mixtures thereof. Whole shale oil, asnoted above, contains distillate fractions with widely differentcompositions and reactivities. Embodiments disclosed hereinadvantageously separate the whole shale oil or partially hydrotreatedwhole shale oil so as to effectively perform the requiredhydroprocessing reactions, namely, diolefins saturation(DOS);hydrodemetallization (HDM); monoolefins saturation (MOS);hydrodenitrogenation (HDN); hydrodesulfurization (HDS);hydrodeoxygenation (HDO); hydrodearomatization (HDA) and gas oilhydrocracking (HYC), without the penalties of severe fouling/pluggingand poor selectivity, i.e., high gas formation, to produce higher-valuedhydrocarbon products.

Referring now to FIG. 1, a simplified flow diagram of an integratedprocess for upgrading crude shale-derived oils, such as those producedby oil shale retorting or in situ extraction or mixtures thereof,according to embodiments herein is illustrated. A whole shale oil 10 maybe fed to a fractionator 12 to fractionate the whole shale oil into afirst fraction 14, including naphtha, kerosene and diesel (NKD), and asecond fraction 16, including gas oil and residuum (AGO).

The first fraction 14 and hydrogen 15 may be fed to a first-stage NKDhydroprocessing reactor 18 containing a hydrogenation catalyst tosaturate diolefins contained in the first fraction. Followinghydrogenation, an effluent 20 may be recovered from the first-stage NKDhydroprocessing reactor 18.

The effluent 20 from the first-stage NKD hydroprocessing reactor 18 maythen be fed, without phase separation, to a second-stage NKDhydroprocessing reactor 22. Second-stage NKD hydroprocessing reactor 22may be operated in an upflow mode and may contain catalysts to performhydrodemetallization and saturation of mono-olefins in the effluent 20recovered from the first-stage NKD hydroprocessing reactor 18. Followingreaction in second-stage NKD hydroprocessing reactor 22, an effluent 24may be recovered from the second-stage NKD hydroprocessing reactor.

The effluent 24 from the second-stage NKD hydroprocessing reactor 22 maythen be fed, without phase separation, to a third-stage NKDhydroprocessing reactor 26. Third-stage NKD hydroprocessing reactor mayinclude one or more beds of catalyst to perform hydrodenitrogenation,hydrodesulfurization, hydrodeoxygenation, and hydrodearomatization ofthe effluent 24 recovered from the second-stage NKD hydroprocessingreactor 22. Following hydroprocessing, an effluent 28 may be recoveredfrom the third-stage NKD hydroprocessing reactor 26.

The second fraction 16 and hydrogen 29 may be fed to a first-stage AGOhydroprocessing reactor 30. First-stage AGO hydroprocessing reactor 30may be operated in an upflow mode and may containing catalysts toperform hydrodemetallization of the second fraction 16. Followingreaction, an effluent 32 may be recovered from the first-stage AGOhydroprocessing reactor 30.

The effluent 32 from the first-stage AGO hydroprocessing reactor 30 maybe fed, without phase separation, to a second-stage AGO hydroprocessingreactor 34. Second-stage AGO hydroprocessing reactor 34 may include oneor more beds of catalyst each containing a catalyst to perform one ormore of hydrotreating and hydrocracking of the effluent 32 recoveredfrom the first-stage AGO hydroprocessing reactor 30. Followinghydrotreating and/or hydrocracking, an effluent 36 may be recovered fromthe second-stage AGO hydroprocessing reactor 34.

The effluents 28, 36 from the second-stage AGO hydroprocessing reactor34 and the third-stage NKD hydroprocessing reactor 26 may be combined toform a mixed stream 38 and fed to a separation train 40. Separationtrain 40 may include one or more distillation and/or extractivedistillation columns useful for separating the effluents into two ormore hydrocarbon fractions. In some embodiments, such as illustrated inFIG. 1, the two or more fractions may include at least one of light gasbyproducts and unreacted hydrogen 42, kerosene 44, diesel 46, and aresiduum fraction 48. Other hydrocarbon fractions may also be recoveredin various embodiments.

The processing of the first fraction 14 and the second fraction 16 maybe performed on a continuous basis. Upflow reactors 22, 30, however, mayrequire more frequent catalyst changes as compared to reactors 18, 26,34. Bypass lines 25, 33 may be provided for bypassing at least one ofthe second-stage NKD hydroprocessing reactor 22 and the first-stage AGOhydroprocessing reactor 30 to replace catalyst within the reactors whilecontinuing to operate the remainder of the process, includingfractionation in columns 12, 40, and reaction in reactors 18, 26, 34.The ability to bypass reactors 22 and 30 may allow replacement of thecatalyst in reactors 22, 30 without shutting down the remainder of theprocess, increasing unit uptime and continued conversion of whole shaleoils to useful hydrocarbons.

Referring now to FIG. 2, a simplified flow diagram of an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting or in situ extraction or mixtures thereof according toembodiments herein is illustrated. A whole shale oil 210 may be fed to afractionator 212 to fractionate the whole shale oil into a firstfraction 214, including naphtha, kerosene and diesel (NKD), and a secondfraction 216, including gas oil and residuum (AGO).

The first fraction 214 and hydrogen 215 may be fed to a first-stage NKDhydroprocessing reactor 218 containing a hydrogenation catalyst tosaturate diolefins contained in the first fraction. Followinghydrogenation, an effluent 220 may be recovered from the first-stage NKDhydroprocessing reactor 218.

The effluent 220 from the first-stage NKD hydroprocessing reactor 218may then be fed, without phase separation, to a second-stage NKDhydroprocessing reactor 222. Second stage NKD hydroprocessing reactor222 may be operated in an upflow mode and may contain catalysts toperform hydrodemetallization and saturation of mono-olefins in theeffluent 220 recovered from the first-stage NKD hydroprocessing reactor218. Following reaction in second stage NKD hydroprocessing reactor 222,an effluent 224 may be recovered from the second-stage NKDhydroprocessing reactor.

The second fraction 216 and hydrogen 227 may be fed to a first-stage AGOhydroprocessing reactor operated in an upflow mode and containingcatalysts to perform hydrodemetallization. Following reaction, aneffluent 228 may be recovered from the first-stage hydroprocessing AGOreactor 226.

The effluent 228 from the first-stage AGO hydroprocessing reactor 226may be fed, without phase separation, to a second-stage AGOhydroprocessing reactor 230 having one or more beds of catalyst eachcontaining a catalyst to perform one or more of hydrotreating andhydrocracking of the effluent from the first-stage AGO hydroprocessingreactor 226. Following hydrotreating and/or hydrocracking, an effluent232 may be recovered from the second-stage AGO hydroprocessing reactor230.

The effluents 224, 232 from the second-stage NKD hydroprocessing reactor222 and the second-stage AGO hydroprocessing reactor 230, respectively,may then be fed, without phase separation, to a third-stagehydroprocessing reactor 234 having one or more beds of catalyst toperform hydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation,and hydrodearomatization of the combined effluent stream. Followinghydroprocessing, an effluent 236 may be recovered from the third-stagehydroprocessing reactor 234.

The effluent 236 from the third-stage hydroprocessing reactor 234 maythen be fed to a separation train 238. Separation train 238 may includeone or more distillation and/or extractive distillation columns usefulfor separating the effluents into two or more hydrocarbon fractions. Insome embodiments, such as illustrated in FIG. 2, the two or morefractions may include at least one of light gas byproducts and unreactedhydrogen 242, kerosene 244, diesel 246, and a residuum fraction 248.Other hydrocarbon fractions may also be recovered in variousembodiments.

The processing of the first fraction 214 and the second fraction 216 maybe performed on a continuous basis. Upflow reactors 222, 226, however,may require more frequent catalyst changes as compared to reactors 218,230, 234. Bypass lines 225, 229 may be provided for bypassing at leastone of the second-stage reactor 222 and the first-stage reactor 226 toreplace catalyst within the reactors while continuing to operate theremainder of the process, including fractionation in columns 212, 238,and reaction in reactors 218, 230, 234. The ability to bypass reactors222 and 226 may allow replacement of the catalyst in reactors 222, 226without shutting down the remainder of the process, increasing unituptime and continued conversion of whole shale oils to usefulhydrocarbons.

As described above with respect to FIGS. 1 and 2, followingfractionation of the whole shale oil, the combined light distillatesstream (including naphtha, kerosene, and diesel-range material) is firstsent through a diolefins saturation (“DOS”) reactor that operates atrelatively low temperatures. The DOS reactor is followed by an upflowreactor (UFR) that is loaded with different types ofhydrodemetallization (HDM) catalysts. Some olefins also get saturated inthis reactor. Following the HDM reactor, the light distillates are sentto a primary hydrotreating reactor where hydrodesulfurization (NDS),hydrodentirogenation (HDN), hydrodeoxygenation (HDO) and some aromaticssaturation reactions occur. The primary hydrotreating reactor effluentis sent to a separator where most of the light gases, ammonia, andhydrogen sulfide are stripped out. The combined heavy distillates stream(including gas oil and resid) is first sent through an upflow reactor(UFR) that is loaded with different types of hydrodemetallization (HDM)catalysts. Olefins may also get saturated in this reactor. Following theHDM reactor, the heavy distillates are sent to a hydroprocessing reactorfor hydrocracking and hydrotreating of the heavy hydrocarbons.

The hydrotreated heavy and light distillate streams may then be furtherprocessed, either together or separately, to recover useful hydrocarbonproducts. For example, the hydrotreated streams may be let down inpressure for removal of ammonia and hydrogen sulfide. Optionally, theresulting liquid streams may also be fed to an aromatics saturationreactor (not illustrated in FIGS. 1 and 2). The reactor effluents areflashed to recover hydrogen and the liquid hydrocarbon products may besent to fractionation for product recovery. When it is desired toproduce syncrude, only a simple stripper will be required for productrecovery. More complex distillation trains may be used to producehydrocarbon products of discrete boiling ranges.

The reaction section of the processes of FIGS. 1 and 2 thus includes twoparallel hydrotreating reactor systems in the same high pressure loop.One reactor system hydrotreats the heavy gasoil feed, while the secondreactor system hydrotreats the combined naphtha, kerosene and diesel toremove contaminants. One alternative scheme, as noted above, includes anaromatics saturation reactor to upgrade the hydrotreated whole shale oilto finished products. Without the aromatics saturation reactor, eitherblending with straight run products to achieve final productspecifications or production of whole crude are also available options.The reaction section may also contain equipment for separation ofhydrogen-rich gas from the last stage reactor effluents. This gas streammay be compressed and recycled back through the high-pressure reactorloop and combined with make-up hydrogen as necessary.

While described above with respect to the simplified flow diagramspresented for FIGS. 1 and 2, various pieces of equipment may be used totreat the feeds and effluents from each reactor and separation system,including heat exchangers, filters, condensers, reboilers, and otherequipment known to those skilled in the art. The following provides amore detailed description of one embodiment of the process. Similaraspects may be extrapolated to other embodiments presented herein, suchas the flow schemes of FIGS. 3-5, described below, which providealternative means for parallel hydrotreating of the heavy gas oil feedand the NKD fraction.

The raw naphtha, kerosene and diesel stream (NKD) may be sent through afeed filter and pumped up to reactor loop pressure by feed pumps. Insome embodiments, the feed filter may be an automatic backwash-typefilter. For shale oil derived feedstock, a 1 micron sock filterfollowing the automatic backwash filter may also be used. The filteredfresh NKD feed is mixed with preheated hydrogen-rich recycle gas andsent through a naphtha, kerosene and diesel Diolefins Saturation Reactor(NKD DOS). The NKD DOS reactor operates at relatively high liquid hourlyspace velocity (LHSV) and a relatively low temperature to facilitatediolefins saturation, and thus preventing the formation of catalystfouling gums through polymerization reactions. The effluent from NKD DOSmay be preheated in a feed effluent exchanger and sent through an NKDUpward Flow Guard Reactor (NKD UFR) where metal contaminants areremoved. The effluents from the NKD UFR are further heated in a secondfeed/effluent exchanger and then in the reactor feed furnace. The inlettemperature to the reactor feed furnace may be controlled by adjustingthe oil feed bypass around the feed/effluent exchangers to maintainsufficient heat input in the furnace. Maintaining this heat input in thefurnace allows stable reactor inlet temperature control and providesoperators with the ability to quickly reduce the reactor inlettemperature in an emergency situation.

From the reactor furnace, the stream enters naphtha, kerosene, dieselHDT Reactor (NKD HDT). The NKD HDT Reactor operates at somewhat lowerLHSV and higher temperatures for HDS and HDN together with olefinssaturation. Here, the oil feed is hydrotreated and partially convertedinto products. The exothermic HDS, HDN and saturation reactions resultin a large heat release that increases the temperature of the reactants.This increased temperature further increases the rate of reaction. Inorder to control this temperature rise and likewise, the rate ofreaction, the catalyst may be separated into multiple beds in the NKDHDT reactor. Cold recycle gas may be introduced between the beds toquench the reacting fluids and thereby control the amount of temperaturerise and the rate of reaction.

In a parallel circuit, the raw gasoil stream is sent through a feedfilter and a 1 micron sock filter and then pumped up to reactor looppressure by feed pumps. The filtered fresh oil feed is mixed withpreheated hydrogen-rich recycle gas, preheated in a feed/effluentexchanger and sent through Gasoil Upward Flow Guard Reactor (Gasoil UFR)where metal contaminants are removed. The effluent from Gasoil UFR isfurther preheated in a second feed/effluent exchanger and then in areactor feed furnace. The inlet temperature to the reactor feed furnacemay be controlled by adjusting the oil feed bypass around thefeed/effluent exchangers to maintain sufficient heat input in thefurnace. Maintaining this heat input in the furnace allows stablereactor inlet temperature control and provides operators with theability to quickly reduce the reactor inlet temperature in an emergencysituation.

From the reactor furnace, the stream enters Gasoil HDT Reactor (GasoilHDT). Here, the oil feed is hydrotreated and partially converted intoproducts. The exothermic hydrocracking and saturation reactions resultin a large heat release that increases the temperature of the reactants.This increased temperature further increases the rate of reaction. Inorder to control this temperature rise and, likewise, the rate ofreaction, the catalyst is separated into multiple beds in the Gasoil HDTreactor. Cold recycle gas may be introduced between the beds to quenchthe reacting fluids and thereby control the amount of temperature riseand the rate of reaction.

Reactor internals between the catalyst beds may also be designed toensure thorough mixing of the reactants with quench gas and gooddistribution of the vapor and liquid flowing down to the next bed. Gooddistribution of the reactants across the catalyst bed prevents hot spotsand maximizes catalyst performance and life. For example, an ISOMIXsystem, available from Chevron Lummus Global, may be used to providecomplete mixing and equilibration of reactants between catalyst beds,correcting any temperature and concentration maldistributions, with lowpressure drop while using minimal reactor volume. ISOMIX allows new andretrofitted reactors to successfully employ new high-activity catalystswith very low risk of reactor temperature maldistribution and hot spotsoften associated with high-activity catalysts.

The Gasoil HDT reactor effluent includes light vaporized hydrocarbons,distillate oils, heavy unconverted oil, and excess hydrogen not consumedin the reaction. The effluent stream is cooled by heat exchange with theoil-gas reactor feed mixture before it is sent through a cascade ofseparators.

The effluents from both NKD HDT and Gasoil HDT reactors are cooled byexchanging heat against their corresponding feeds and enter separate HotHigh Pressure Separators (HHPS) that operate in parallel. The vaporstreams from the respective HHPS are combined and cooled by heatexchange with the reactor feed gas and a Cold Low Pressure Separator(CLPS) liquid. At this point, water may be continuously injected intothe inlet piping of the HHPS vapor air cooler to prevent the depositionof salts in the air cooler tubes. Without the water injection, ammonia(NH3) and hydrogen sulfide (H2S), which are formed in the reactor byhydrodesulfurization of sulfur species and hydrodenitrogenation ofnitrogen species in the feed, can form solid ammonium bisulfide (NH4HS)at cooler temperatures. This solid can deposit on the air cooler tubes,reduce heat transfer, and eventually plug the tubes. Since ammoniumbisulfide is soluble in water, the continuous presence of water willdissolve any NH4HS that has formed and thus prevent the deposition ofNH4HS solids in the air cooler tubes. Both the water injection pipingand the air cooler inlet piping must meet design specifications toensure uniform distribution of water to the effluent air cooler. Thisunit is designed to reuse condensate from the fractionator overhead.

The HHPS effluent vapor air cooler cools the effluent for maximumrecovery of hydrocarbon liquids from the vapor. The cooled effluent isseparated into its hydrogen rich vapor, hydrocarbon liquid, and waterphases. The sour water stream containing ammonium bisulfide is sent tosour water stripping. The hydrocarbon liquid is fed to the CLPS.

The hydrogen-rich gas from the CHPS flows into a knockout drum before itenters an H2S absorber. Because of the extremely high H2S concentration,a high-pressure Amine Absorber may be used for maintaining recycle gasquality. For a required hydrogen partial pressure, the high pressureloop operating pressure can be significantly lowered by increasing therecycle gas purity using an amine absorber. Such may also be beneficialfor aromatics separation and facilitate use of low reactor temperatures.

The sweetened gas then flows to a knockout drum and then to the recyclegas compressor. The compressor suction line is heat traced to ensure aliquid-free vapor. The recycle compressor delivers the recycle gas tothe reactor loop. There is a purge line located upstream of the recyclegas compressor that can be used if necessary to send amine-sweetenedrecycle gas to the flare. However, no high pressure purge gas rate isrequired in normal operation. An emergency dump line may be locatedupstream of the recycle gas compressor to allow quick reduction ofpressure in the recycle loop, if necessary, to control reactortemperatures during a loss of the recycle compressor or other upsets.

Part of the recycle compressor discharge gas is routed to the reactorsas quench to control the reactor temperature. The remaining recycle gasthat is not used as quench is combined with make-up hydrogen to becomethe reactor feed gas. Reliable, uninterrupted operation of the recyclecompressor may facilitate safe operation of the plant. One type ofreliable recycle compressor is a centrifugal machine with a steamturbine driver. The reactor feed gas for both stages is heated byexchange with the HHPS vapor before combining with the oil feed streamsto each reaction stage.

The process requires a continuous supply of high-pressure make-uphydrogen. In addition to chemical consumption, hydrogen leaves thesystem as offgas from the cold low pressure separator (CLPS) asdissolved hydrogen in the product distillation feed and may also be lostthrough system leaks.

In the high-pressure circuit, the liquid stream from the HHPS is reducedin pressure and fed to the Hot Low Pressure Separator (HLPS). The liquidfrom the Gasoil HLPS is routed directly to the product stripper and thevapor stream is routed to an air cooler and further to the Gasoil ColdLow Pressure Separator (CLPS). In the Gasoil CLPS, the liquid isseparated from the vapor and is heated by heat exchange with the GasoilHHPS vapor before combining with liquid from the NKD HHPS and enteringNKD HLPS. The sour gas from the gasoil CLPS and NKD HLPS may be sent forfurther treatment and hydrogen recovery. The liquid from NKD HLPS ispumped up to the Naphtha/Kerosene/Diesel/Aromatics Saturation (NKD ASAT)reactor loop pressure by the feed pumps. It is preheated in a feedeffluent exchanger, mixed with preheated hydrogen rich recycle-gas. Themixture is further preheated against NKD effluent and fed to NKD ASATreactor.

The preheated stream enters the NKD ASAT reactor. The catalyst in thisreactor promotes aromatics saturation and further hydrodesulfurization.Cold recycle gas is introduced between the beds to quench the reactingfluids and thereby control the amount of temperature rise and the rateof reaction.

The effluents from both NKD ASAT reactors is cooled by exchanging heatwith NKD ASAT reactor feed, recycle hydrogen rich gas, and NKD productstripper feed, and then sent through an effluent aircooler. At thispoint, water is continuously injected into the inlet piping of the NKDASAT reactor effluent air cooler to prevent the deposition of salts inthe air cooler tubes.

The NKD ASAT reactor effluent air cooler cools the effluent for maximumrecovery of hydrocarbon liquids from the vapor. The cooled effluent isseparated into its hydrogen rich vapor, hydrocarbon liquid, and waterphases in the NKD CHPS. The vapor joins vapor from the gasoil CHPS andis sent to a knockout drum before it enters an H2S absorber. The sourwater stream containing ammonium bisulfide is sent to sour waterstripping. The hydrocarbon liquid is fed to the NKD CLPS.

In the NKD CLPS, the liquid is separated from the vapor and is heated byheat exchange with NKD ASAT reactor effluent before being fed to the NKDProduct Stripper. The sour gas from the NKD CLPS is sent for furthertreatment and hydrogen recovery.

When producing synthetic crude, the entire aromatics saturation step canbe eliminated and the hydrogen and product recovery sections combined.

The fractionation section may include a Gasoil product stripper, aNaphtha/Kerosene/Diesel Product Stripper, and a Product Fractionator.The fractionation section may be designed to separate reaction productsinto Light Ends for separation of LPG, Naphtha, Kerosene, Diesel, andtreated Gasoil.

The product stripper's primary function is to separate light products atsufficient pressure to feed a deethanizer column in the light endsrecovery section without the need for a sour gas compressor.

After flashing hydrogen gas, the remaining liquid reactor effluent isrouted to the product stripper. In the product stripper, the gas,propane, butane, and some unstabilized naphtha are removed from thereactor effluent for processing in the light ends recovery section. Theheavier products of Gasoil product stripper are then routed from thebottom to tankage is a treated gasoil. The heavier products of theNaphtha/Kerosene/Diesel Product Stripper, a combination of treatednaphtha, jet fuel and diesel, are routed from the bottom to thefractionator feed furnace for heating before entering the fractionator,which operates at low pressure.

The product stripper is refluxed with unstabilized naphtha and strippedwith superheated steam. Water from the product stripper cannot berecycled back to injection water system in the reactor because of thehigh concentration of ammonium bisulfide.

The fractionator system separates the NKD ASAT reactor effluent intonaphtha, kerosene, and diesel. Overhead vapor from the fractionator iscondensed in the total condenser and sent to the overhead accumulator. Afixed amount of the liquid is returned to the fractionator as reflux. Avapor line from the overhead accumulator to the flare system is includedshould any vapors build up, The net liquid product from the accumulatoris naphtha, which is sent to the light ends recovery section. Water fromthe overhead accumulator is sent to the water injection drum in thereaction section.

Trays are provided in the fractionator for separation of naphtha andkerosene and between kerosene and diesel. A liquid draw is taken fromthe fractionator and stripped in the kerosene sidecut stripper. Theoverhead vapor is returned to the fractionator. The stripper bottomsstream is pumped by the kerosene product pump and then cooled. Thecooled product is sent to tankage. The fractionator bottoms stream ispumped by the fractionator bottoms pump to the fractionator feed preheatexchanger and after cooling sent to tankage.

As described above, catalysts may be provided in each of the reactors toperform various hydroprocessing operations including hydrogenation(diolefins saturation, monoolefins saturation and/or aromaticssaturation), hydrodeoxygenation, hydrodemetallization,hydrodenitrogenation, hydrocracking, hydrodesulfurization andhydrotreating. A hydrotreating catalyst, for example, may include anycatalyst composition that may be used to catalyze the hydrogenation ofhydrocarbon feedstocks to increase its hydrogen content and/or removeheteroatom contaminants. A hydrocracking catalyst, for example, mayinclude any catalyst composition that may be used to catalyze theaddition of hydrogen to large or complex hydrocarbon molecules as wellas the cracking of the molecules to obtain smaller, lower molecularweight molecules.

Hydroprocessing catalyst compositions for use in processes according toembodiments disclosed herein are well known to those skilled in the artand several are commercially available from W.R. Grace & Co., CriterionCatalysts & Technologies, and Albemarle, among others. Suitablehydroconversion catalysts may include one or more elements selected fromGroups 4-12 of the Periodic Table of the Elements. In some embodiments,hydroconversion catalysts according to embodiments disclosed herein maycomprise, consist of, or consist essentially of one or more of nickel,cobalt, tungsten, molybdenum and combinations thereof, eitherunsupported or supported on a porous substrate such as silica, alumina,titania, or combinations thereof. As supplied from a manufacturer or asresulting from a regeneration process, the hydroconversion catalysts maybe in the form of metal oxides, metal hydrides, or metal sulfides, forexample. In some embodiments, the catalysts may be pre-sulfided and/orpre-conditioned prior to introduction to the reactor(s).

Hydrotreating and hydrogenation catalysts that may be useful includecatalysts generally composed of a hydrogenation component, selected fromGroup 6 elements (such as molybdenum and/or tungsten) and Group 8-10elements (such as cobalt and/or nickel), or a mixture thereof, which maybe supported on an alumina support. Phosphorous (Group 15) oxide isoptionally present as an active ingredient. A typical catalyst maycontain from 3 to 35 wt % hydrogenation components, with an aluminabinder. The catalyst pellets may range in size from 1/32 inch to ⅛ inch,and may be of a spherical, extruded, trilobate or quadrilobate shapeCatalyst layer(s) for demetallization, when present, may comprisecatalyst(s) having an average pore size ranging from 125 to 225Angstroms and a pore volume ranging from 0.5-1.1 cm³/g. Catalystlayer(s) for denitrification/desulfurization may comprise catalyst(s)having an average pore size ranging from 100 to 190 Angstroms with apore volume of 0.5-1.1 cm³/g. U.S. Pat. No. 4,990,243 describes ahydrotreating catalyst having a pore size of at least about 60Angstroms, and preferably from about 75 Angstroms to about 120Angstroms. A demetallation catalyst useful for the present process isdescribed, for example, in U.S. Pat. No. 4,976,848, the entiredisclosure of which is incorporated herein by reference for allpurposes. Likewise, catalysts useful for desulfurization of heavystreams are described, for example, in U.S. Pat. Nos. 5,215,955 and5,177,047, the entire disclosures of which are incorporated herein byreference for all purposes. Catalysts useful for desulfurization ofmiddle distillate, vacuum gas oil streams and naphtha streams aredescribed, for example, in U.S. Pat. No. 4,990,243, the entiredisclosures of which are incorporated herein by reference for allpurposes.

Reactors for hydrogenating the light distillate fraction, such asreactors 18 and 218, may be operated at reaction conditions including atemperature in the range from about 100° C. to about 250° C., a hydrogenpartial pressure in the range from about 400 to about 500 psi, and aliquid hourly space velocity in the range from about 2 to about 6 L perhour per L catalyst.

Reactors for hydrodemetallization of the light distillate fraction, suchas reactors 22 and 222, may be operated at reaction conditions includinga temperature in the range from about 200° C. to about 440° C., ahydrogen partial pressure in the range from about 400 to about 2600 psi,and a liquid hourly space velocity in the range from about 0.5 to about5.0 L per hour per L catalyst.

Reactors for hydrotreating (HDN, HDS, HDO, HDA, etc.) the lightdistillate fraction, such as reactor 26, may be operated at reactionconditions including a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst.

Reactors for hydrodemetallization of the gas oil fraction, such asreactors 30, 226, may be operated at reaction conditions including atemperature in the range from about 200° C. to about 440° C., a hydrogenpartial pressure in the range from about 400 to about 2600 psi, and aliquid hourly space velocity in the range from about 0.5 to about 5.0 Lper hour per L catalyst.

Reactors for hydrotreating (HDN, HDS, HDO, HDA, etc.) of the gas oilfraction, such as reactors 34, 230, may be operated at reactionconditions including a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst.

Reactors for hydrotreating HDS, HDO, HDA, etc.) a mixture of the lightdistillate fraction and the heavy distillate fraction, such as reactor234, may be operated at reaction conditions including a temperature inthe range from about 280° C. to about 440° C., a hydrogen partialpressure in the range from about 800 to about 2600 psi, and a liquidhourly space velocity in the range from about 0.3 to about 4.0 L perhour per L catalyst.

In reactor 18 carrying out the diolefins saturation reactions (DOS), thecatalysts used are tailored for double bond saturation and littleheteroatom removal (HDN, HDS, HDO, HDA) is achieved. The diolefinssaturation is in the range from about 90% to about 100%.

In reactors carrying out the hydrodemetallization reactions, HDN removalis in the range from about less than 1% to about 15%. HDS removal is inthe range from about less than 1% to about 15%. HDO removal is in therange from about 10% to about 50%. HDM removal is in the range fromabout 70% to about 100%. These include reactors 22, 30.

In reactors carrying out the hydrotreating reactions, the HDN removal isin the range from about 40% to about 100%. The HDS removal is in therange from about 40% to about 100%. The HDO removal is in the range fromabout 80% to about 100%. The HDM removal is in the range from about 30%to about 100%. These include reactors 26, 34.

The processes as described with respect to FIGS. 1 and 2 may includeadditional hydrocarbon feedstocks in addition to whole shale oil. Forexample, one or more additional hydrocarbon feedstocks may be fed tofractionator 12, 212. The one or more additional hydrocarbon feedstocksmay include hydrocarbonaceous materials derived from thermal tars,bitumen, coke oven tars, asphaltenics, coal gasification tars,biomass-derived tars, black liquor tars, or reactive hydrocarbonaceousmaterials derived from thermal tars, bitumen, coke oven tars,asphaltenics, coal gasification tars, biomass-derived tars, and blackliquor tars, produced in one or more of thermal cracking, pyrolysis, andretorting processes. As another example, one or more additionalhydrocarbon feedstocks may be fed to hydroprocessing reactors 30, 226.

Referring now to FIG. 3, a simplified flow diagram of an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting according to embodiments herein is illustrated. A whole shaleoil 310 and hydrogen 311 may be contacted in a first-stagehydroprocessing reactor 312 containing hydrogenation catalysts tosaturate diolefins contained in the whole shale oil. Followinghydrogenation, an effluent 314 may be recovered from the first-stagehydroprocessing reactor 312.

The effluent 314 from the first-stage hydroprocessing reactor 312 may befed, without phase separation, to a second-stage hydroprocessing reactor316, which may be operated in an upflow mode and containing catalysts toperform hydrodemetallization and saturation of mono-olefins in theeffluent 314 from the first-stage hydroprocessing reactor 316. Followinghydroprocessing, an effluent 318 may be recovered from the second-stagehydroprocessing reactor 316.

Effluent 318 may then be fed to a fractionation system 320 to separateeffluent 318 into a first partially hydrotreated fraction 322, includingnaphtha, kerosene and diesel, and a second partially hydrotreatedfraction 324, including gas oil and resid.

The first partially hydrotreated fraction 322 may then be fed to athird-stage hydroprocessing reactor 326 having one or more beds ofcatalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the first partiallyhydrotreated fraction 322. Following hydroprocessing, an effluent 328may be recovered from the third-stage hydroprocessing reactor 326.

The second partially hydrotreated fraction 324 may also be fed to athird-stage hydroprocessing reactor 330 having one or more beds ofcatalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the second partiallyhydrotreated fraction 324. Following hydroprocessing, an effluent 332may be recovered from the third-stage hydroprocessing reactor 330.

Unreacted hydrogen may be recovered from the top of fractionation system320 along with first partially hydrotreated fraction 322. Additionalhydrogen may be fed, if necessary, via flow line 323 for hydroprocessingin third-stage reactor 326. Likewise, dissolved hydrogen may becontained in the second partially hydrotreated fraction 324. Ifnecessary, additional hydrogen may be fed via flow line 325 forhydroprocessing in third-stage reactor 330.

The effluents 328, 332 from third-stage hydroprocessing reactors 326,330, respectively, may then be fed to a separation train 334 to recovertwo or more hydrocarbon fractions. Separation train 334 may include oneor more distillation and/or extractive distillation columns useful forseparating the effluents into two or more hydrocarbon fractions. In someembodiments, such as illustrated in FIG. 3, the two or more fractionsmay include at least one of light gas byproducts and unreacted hydrogen342, kerosene 344, diesel 346, and a residuum fraction 348. Otherhydrocarbon fractions may also be recovered in various embodiments.

Referring now to FIG. 4, a simplified flow diagram of an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting according to embodiments herein is illustrated. A whole shaleoil 410 and hydrogen 411 may be contacted in a first-stagehydroprocessing reactor 412 containing hydrogenation catalysts tosaturate diolefins contained in the whole shale oil. Followinghydrogenation, an effluent 414 may be recovered from the first-stagehydroprocessing reactor 412.

The effluent 414 from the first-stage hydroprocessing reactor 412 may befed, without phase separation, to a second-stage hydroprocessing reactor416, which may be operated in an upflow mode and containing catalysts toperform hydrodemetallization and saturation of mono-olefins in theeffluent 414 from the first-stage hydroprocessing reactor 412. Followinghydroprocessing, an effluent 418 may be recovered from the second-stagehydroprocessing reactor 416.

Effluent 418 may then be fed to a fractionation system 420 to separateeffluent 418 into a first partially hydrotreated fraction 422, includingnaphtha, kerosene and diesel, and a second partially hydrotreatedfraction 424, including gas oil and resid.

The second partially hydrotreated fraction 424 may be fed to athird-stage hydroprocessing reactor 430 having one or more beds ofcatalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the second partiallyhydrotreated fraction 424. Following hydroprocessing, an effluent 432may be recovered from the third-stage hydroprocessing reactor 430.

The first partially hydrotreated fraction 422 and effluent 432 may thenbe fed to a third-stage hydroprocessing reactor 426 having one or morebeds of catalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the first partiallyhydrotreated fraction 422 and effluent 432. Following hydroprocessing,an effluent 428 may be recovered from the third-stage hydroprocessingreactor 426.

Unreacted hydrogen may be recovered from the top of fractionation system420 along with first partially hydrotreated fraction 422. Additionalhydrogen may be fed, if necessary, via flow line 423 for hydroprocessingin third-stage reactor 426. Likewise, dissolved hydrogen may becontained in the second partially hydrotreated fraction 424. Ifnecessary, additional hydrogen may be fed via flow line 425 forhydroprocessing in third-stage reactor 430.

The effluent 428 from third-stage hydroprocessing reactors 426 may thenbe fed to a separation train 434 to recover two or more hydrocarbonfractions. Separation train 434 may include one or more distillationand/or extractive distillation columns useful for separating theeffluents into two or more hydrocarbon fractions. In some embodiments,such as illustrated in FIG. 4, the two or more fractions may include atleast one of light gas byproducts and unreacted hydrogen 442, kerosene444, diesel 446, and a residuum fraction 448. Other hydrocarbonfractions may also be recovered in various embodiments.

Referring now to FIG. 5, a simplified flow diagram of an integratedprocess for upgrading crude shale-derived oils produced by oil shaleretorting according to embodiments herein is illustrated. A whole shaleoil 510 and hydrogen 511 may be contacted in a first-stagehydroprocessing reactor 518 containing catalysts to saturate diolefinscontained in the whole shale oil. Following hydrogenation, an effluent514 may be recovered from the first-stage hydroprocessing reactor 518.

The effluent 514 from the first-stage hydroprocessing reactor 518 may befed, without phase separation, to a second-stage hydroprocessing reactor522, which may be operated in an upflow mode and containing catalysts toperform hydrodemetallization and saturation of mono-olefins in theeffluent 514 from the first-stage hydroprocessing reactor 522.Additional hydrogen 515 may be added as necessary. Followinghydroprocessing, an effluent 524 may be recovered from the second-stagehydroprocessing reactor 522.

The effluent 524 from the second-stage hydroprocessing reactor 522 maybe fed, with or without phase separation, to a third-stagehydroprocessing reactor 526 having one or more beds of catalyst toperform hydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation,and aromatics saturation of the effluent from the second stagehydroprocessing reactor 522. Following hydroprocessing, an effluent 528may be recovered from the third-stage hydroprocessing reactor 526.

The effluent 528 from the third-stage hydroprocessing reactor 526 maythen be fed to a separation train 538. Separation train 538 may includeone or more distillation and/or extractive distillation columns usefulfor separating the effluents into two or more hydrocarbon fractions. Insome embodiments, such as illustrated in FIG. 5, the separation train538 may include an atmospheric distillation column 539 and a vacuumdistillation column 540 to separate the effluent 528 into distillatefractions including light gas fraction 542, naphtha 544, kerosene, jet,and diesel fraction 546, as well as a hydrotreated vacuum gas oilfraction 548 and a low sulfur fuel oil fraction 550.

The hydrotreated vacuum gas oil fraction 548 and hydrogen 549 may thenbe fed to a fourth-stage hydroprocessing reactor 534, which may includeone or more beds of catalyst to perform hydrocracking of thehydrotreated vacuum gas oil fraction. Following hydrocracking, aneffluent 536 may be recovered from the fourth-stage hydroprocessingreactor 534. The effluent 536 from the fourth-stage hydroprocessingreactor 534 may then be fed to the fractionation train 538 forseparation as described above.

The processing of the oil through the reactors illustrated in FIG. 5 maybe performed on a continuous basis. Similar to reactors 222, 226,reactor 522 may require more frequent catalyst changes as compared toreactors 518, 526, 534. Bypass line 555 may be provided for bypassingthe second stage reactor 522 to replace catalyst within the reactorwhile continuing to operate the remainder of the process.

In the embodiment illustrated in FIG. 5, the shale oil feed is firstprocessed in reactors 518, 522, and 526 to achieve near completedemetallation, hydrodenitrification, hydrodesulfurization, and somearomatic saturation. Limited hydrocracking may also result. Followingproduct separation, the vacuum gas oil fraction 548 may be completelyhydrocracked to products by operating at about 60% per pass conversion.Further, the unconverted oil recovered from the vacuum column bottoms isan excellent quality FCC feedstock.

The process of FIG. 5, as well as other embodiments herein, may providefor the capability to adjust the first stage reactor severity to addressfeed variations, may produce high quality FCC feedstocks, and mayprovide a consistent feed to the hydrocracking reactor 534 to producehigh quality mid distillates and ensure extended catalyst life, amongother advantages.

FIGS. 3, 4, and 5 thus illustrate alternative process flow schemes forupgrading the heavy and light components in whole shale oil viahydroprocessing. Catalysts useful in the reactors described with respectto FIGS. 3-5 are similar to those as described above with respect toFIGS. 1 and 2.

Reactors for hydrodemetallization of the whole shale oil, such asreactors 316, 416, and 522 may be operated at reaction conditionsincluding a temperature in the range from about 200° C. to about 400°C., a hydrogen partial pressure in the range from about 400 to about2600 psi, and a liquid hourly space velocity in the range from about 0.5to about 5.0.

Reactors for hydrotreating (HDN, HDS, HDO, ASAT, etc.) the lightdistillate fraction, such as reactor 326, may be operated at reactionconditions including a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst.

Reactors for hydrotreating (HDN, HDS, HDO, ASAT, etc.) a mixture of thelight and heavy distillate, such as reactor 426 and 526, may be operatedat reaction conditions including a temperature in the range from about280° C. to about 440° C., a hydrogen partial pressure in the range fromabout 800 to about 2600 psi, and a liquid hourly space velocity in therange from about 0.3 to about 4.0 L per hour per L catalyst.

Reactors for hydrotreating (HDN, HDS, HDO, ASAT, etc.) the heavydistillate fraction, such as reactors 330, 430 may be operated atreaction conditions including a temperature in the range from about 280°C. to about 440° C., a hydrogen partial pressure in the range from about800 to about 2600 psi, and a liquid hourly space velocity in the rangefrom about 0.3 to about 4.0 L per hour per L catalyst.

In reactors carrying out the diolefins saturation reactions (DOS), thecatalysts used are tailored to double bond saturation and littleheteroatom removal (HDN, HDS, HDO, ASAT) is achieved. The diolefinssaturation is in the range from about 90% to about 100%. These includereactors 312, 412, 518.

In reactors carrying out the hydrodemetallization reactions, HDN removalis in the range from about less than 1% to about 15%. HDS removal is inthe range from about less than 1% to about 15%. HDO removal is in therange from about 10% to about 50%. HDM removal is in the range fromabout 70% to about 100%. These include reactors 226, 316, 416, 522.

In reactors carrying out the hydrotreating reactions, the HDN removal isin the range from about 40% to about 100%. The HDS removal is in therange from about 40% to about 100%. The HDO removal is in the range fromabout 80% to about 100%. The HDM removal is in the range from about 30%to about 100%. These include reactors 230, 234, 326, 426, 526.

In reactor 526, carrying out hydrocracking of the partially hydrotreatedgas oil fraction, the per pass conversion may be in the range from about40% to about 70%.

The processes as described with respect to FIGS. 3-5 may includeadditional hydrocarbon feedstocks in addition to whole shale oil. Forexample, one or more additional hydrocarbon feedstocks may be fed toreactors 312, 412. The one or more additional hydrocarbon feedstocks mayinclude hydrocarbonaceous materials derived from thermal tars; bitumen;coke oven tars; asphaltenics; coal gasification tars; biomass-derivedtars; black liquor tars, such as reactive hydrocarbonaceous materialsfrom one or more of thermal cracking, pyrolysis, and retorting processesderived from thermal tars; bitumen; coke oven tars; asphaltenics; coalgasification tars; biomass-derived tars; black liquor tars. As anotherexample, one or more additional hydrocarbon feedstocks may be fed tohydroprocessing reactors 330, 430, or 534.

As described above, embodiments disclosed herein provide for asplit-flow processing scheme to upgrade whole shale oil. The split flowconcepts described herein, i.e., naphtha and kerosene hydrotreating inone or more stages and gas oil hydrotreating in one or more stages,requires additional equipment as compared to the alternative approach ofwhole oil hydrotreating. While contrary to conventional wisdom asrequiring more capital equipment to achieve the same final productspecifications, the operating efficiency vis a vis on-stream timeefficiency and product quality resulting from the split flow concept farexceed in value the somewhat incrementally higher capital expenditurecosts.

Embodiments herein advantageously eliminate the disadvantages of thealternative whole oil hydroprocessing concept as noted above. Further,the split-processing embodiments disclosed herein optimize catalystutilization, product yields, and hydrogen consumption. These advantagesdirectly translate to lower investment, utilities and manpower. Thedisclosed embodiments avoid numerous problems characteristic to theprevious whole oil hydrotreating solutions, including: inhibiting heavyoil hydrocracking (directly or via reduction of hydrogen partialpressure); cracking away some desired products (diesel) because wholeoil hydrotreating conditions are more severe than diesel hydrotreating;consumption of more hydrogen because of excessive hydrocracking andoversaturation of heavy oils; and impeding of the diesel hydrotreatingfunction by the presence of very large molecules from heavy oil streams.This aspect becomes very important when producing ULSD (ultra-low sulfurdiesel).

While the disclosure includes a limited number of embodiments, thoseskilled in the art, having benefit of this disclosure, will appreciatethat other embodiments may be devised which do not depart from the scopeof the present disclosure. Accordingly, the scope should be limited onlyby the attached claims.

What is claimed:
 1. An integrated process for upgrading crudeshale-derived oils produced by oil shale retorting or by in situextraction or by mixtures thereof that includes the following steps: a.fractionating the whole shale oil into a first fraction comprisingnaphtha, kerosene and diesel and an atmospheric bottoms fractioncomprising gas oil and residuum; b. contacting the first fraction andhydrogen in a first-stage hydroprocessing reactor containing ahydrogenation catalyst to saturate diolefins contained in the firstfraction and recovering an effluent from the first-stage hydroprocessingreactor; c. feeding the effluent from the first-stage hydroprocessingreactor of step (b) without phase separation to a second-stagehydroprocessing reactor operated in an upflow mode and containingcatalysts to perform hydrodemetallization and saturation of mono-olefinsin the effluent from the first-stage hydroprocessing reactor andrecovering an effluent from the second-stage hydroprocessing reactor; d.feeding the effluent from the second-stage hydroprocessing reactor ofstep (c) without phase separation to a third-stage hydroprocessingreactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the effluent from the second-stagehydroprocessing reactor and recovering an effluent from the third-stagehydroprocessing reactor; e. feeding the atmospheric bottoms fraction andhydrogen to a fourth-stage hydroprocessing reactor operated in an upflowmode and containing catalysts to perform hydrodemetallization of theatmospheric bottoms fraction and recovering an effluent from thefourth-stage reactor; f. feeding the effluent from the fourth-stagehydroprocessing reactor of step (e) without phase separation to afifth-stage hydroprocessing reactor having one or more beds of catalysteach containing a catalyst to perform one or more of hydrotreating andhydrocracking of the effluent from the fourth-stage hydroprocessingreactor and recovering an effluent from the fifth-stage hydroprocessingreactor; g. processing the effluents from the fifth-stagehydroprocessing reactor of step (f) and the third-stage hydroprocessingreactor of step (d) in a separation train to recover two or morehydrocarbon fractions.
 2. The process of claim 1, wherein the two ormore hydrocarbon fractions include at least one of naphtha, kerosene,diesel, and a residuum fraction.
 3. The process of claim 1, wherein thecatalyst in the first stage reactor comprises an alumina base extrudedcatalyst with nickel and molybdenum as active metals; the catalyst inthe second stage reactor comprises an alumina base spheroidal catalystwith nickel and molybdenum as active metals; the catalyst in the thirdstage reactor comprises a layered catalyst comprised of a layer ofamorphous base metal Type II extruded catalyst with an organic compoundand nickel and molybdenum as active metals and a layer of base metalextruded catalyst containing both amorphous and zeolitic components withnickel and tungsten as active metals; the catalyst in the fourth stagereactor comprises an alumina base spheroidal catalyst with nickel andmolybdenum as active metals; and the catalyst in the fifth stage reactorcomprises an amorphous base metal Type II extruded catalyst with anorganic compound and nickel and molybdenum as active metals.
 4. Theprocess of claim 1, further comprising: operating the first stagereactor at reaction conditions comprising a temperature in the rangefrom about 100° C. to about 250° C., a hydrogen partial pressure in therange from about 400 psi to about 500 psi, and a liquid hourly spacevelocity in the range from about 2 to about 6 L per hour per L catalyst;operating the second stage reactor at reaction conditions comprising atemperature in the range from about 200° C. to about 440° C., a hydrogenpartial pressure in the range from about 400 to about 2600 psi, and aliquid hourly space velocity in the range from about 0.5 to about 5 Lper hour per L catalyst; operating the third stage reactor at reactionconditions comprising a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst; operating the fourthstage reactor at reaction conditions comprising a temperature in therange from about 200° C. to about 440° C., a hydrogen partial pressurein the range from about 400 to about 2600 PSI, and a liquid hourly spacevelocity in the range from about 0.5 to about 5.0 L per hour per Lcatalyst; operating the fifth stage reactor at reaction conditionscomprising a temperature in the range from about 280° C. to about 440°C., a hydrogen partial pressure in the range from about 800 to about2600 psi, and a liquid hourly space velocity in the range from about 0.3to about 4.0 L per hour per L catalyst.
 5. The process of claim 1,further comprising bypassing at least one of the second-stagehydroprocessing reactor of step (c) and the fourth-stage hydroprocessingreactor of step (e) to replace catalyst within the reactors whilecontinuing to perform steps (a), (b), (d), (f), and (g).
 6. The processof claim 1, further comprising feeding one or more additionalhydrocarbon feedstocks to the fractionation step (a), the one or moreadditional hydrocarbon feedstocks comprising hydrocarbonaceous materialsderived from thermal tars; bitumen; coke oven tars; asphaltenics; coalgasification tars; biomass-derived tars; black liquor tars.
 7. Theprocess of claim 1, further comprising feeding one or more additionalhydrocarbon feedstocks to the fourth-stage hydroprocessing reactor ofcontacting step (e), the one or more additional hydrocarbon feedstockscomprising hydrocarbonaceous materials derived from thermal tars;bitumen; coke oven tars; asphaltenics; coal gasification tars;biomass-derived tars; black liquor tars.
 8. An integrated process forupgrading crude shale-derived oils produced by oil shale retorting or byin situ extraction or combinations thereof that includes the followingsteps: a. fractionating the whole shale oil into a first fractioncomprising naphtha, kerosene and diesel and an atmospheric bottomsfraction comprising gas oil and residuum; b. feeding the first fractionand hydrogen to a first-stage hydroprocessing reactor containing ahydrogenation catalyst to saturate diolefins contained in the firstfraction and recovering an effluent from the first-stage hydroprocessingreactor; c. feeding the effluent from the first-stage hydroprocessingreactor of step (b) without phase separation to a second-stagehydroprocessing reactor operated in an upflow mode and containingcatalysts to perform hydrodemetallization and saturation of mono-olefinsin the effluent from the first-stage hydroprocessing reactor andrecovering an effluent from the second-stage hydroprocessing reactor; d.feeding the atmospheric bottoms fraction and hydrogen to a third-stagehydroprocessing reactor operated in an upflow mode and containingcatalysts to perform hydrodemetallization and recovering an effluentfrom the third-stage hydroprocessing reactor; e. feeding the effluentfrom the third-stage hydroprocessing reactor of step (d) without phaseseparation to a fourth-stage hydroprocessing reactor having one or morebeds of catalyst each containing a catalyst to perform one or more ofhydrotreating and hydrocracking of the effluent from the third-stagehydroprocessing reactor and recovering an effluent from the fourth-stagehydroprocessing reactor; f. feeding the effluent from the second-stagehydroprocessing reactor of step (c) and the effluent from thefourth-stage hydroprocessing reactor of step (e), without phaseseparation, to a fifth-stage hydroprocessing reactor having one or morebeds of catalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the effluents from thesecond- and fourth-stage hydroprocessing reactors and recovering aneffluent from the fifth-stage hydroprocessing reactor; g. processing theeffluents from the fifth-stage hydroprocessing reactor of step f) in aseparation train to recover two or more hydrocarbon fractions.
 9. Theprocess of claim 8, wherein the two or more hydrocarbon fractionsinclude at least one of naphtha, kerosene, diesel, and a residuumfraction.
 10. The process of claim 8, wherein the catalyst in the firststage reactor comprises an alumina base extruded catalyst with nickeland molybdenum as active metals; the catalyst in the second stagereactor comprises an alumina base spheroidal catalyst with nickel andmolybdenum as active metals; the catalyst in the third stage reactorcomprises an alumina base spheroidal catalyst with nickel and molybdenumas active metals; the catalyst in the fourth stage reactor comprises anamorphous base metal Type II extruded catalyst with an organic compoundand nickel and molybdenum as active metals; and the catalyst in thefifth stage reactor comprises a layered catalyst comprised of a layer ofamorphous base metal Type II extruded catalyst with an organic compoundand nickel and molybdenum as active metals and a layer of base metalextruded catalyst containing both amorphous and zeolitic components withnickel and tungsten as active metals.
 11. The process of claim 8,further comprising: operating the first stage reactor at reactionconditions comprising a temperature in the range from about 100° C. toabout 250° C., a hydrogen partial pressure in the range from about 400to about 500 psi, and a liquid hourly space velocity in the range fromabout 2 to about 6 L per hour per L catalyst; operating the second stagereactor at reaction conditions comprising a temperature in the rangefrom about 200° C. to about 400° C., a hydrogen partial pressure in therange from about 400 to about 2600 psi, and a liquid hourly spacevelocity in the range from about 0.5 to about 5.0 L per hour per Lcatalyst; operating the third stage reactor at reaction conditionscomprising a temperature in the range from about 280° C. to about 440°C., a hydrogen partial pressure in the range from about 800 to about2600 psi, and a liquid hourly space velocity in the range from about 0.3to about 4.0 L per hour per L catalyst; operating the fourth stagereactor at reaction conditions comprising a temperature in the rangefrom about 280° C. to about 440° C., a hydrogen partial pressure in therange from about 800 to about 2600 psi, and a liquid hourly spacevelocity in the range from about 0.3 to about 4.0 L per hour per Lcatalyst; operating the fifth stage reactor at reaction conditionscomprising a temperature in the range from about 280° C. to about 440°C., a hydrogen partial pressure in the range from about 800 to about2600 psi, and a liquid hourly space velocity in the range from about 0.3to about 4.0 L per hour per L catalyst.
 12. The process of claim 8,further comprising bypassing at least one of the second stage reactor ofstep (c) and the fourth stage reactor of step (d) to replace catalystwithin the reactors while continuing to perform steps (a), (b), (e),(f), and (g).
 13. The process of claim 8, further comprising feeding oneor more additional hydrocarbon feedstocks to the fractionation step (a),the one or more additional hydrocarbon feedstocks comprisinghydrocarbonaceous materials derived from thermal tars; bitumen; cokeoven tars; asphaltenics; coal gasification tars; biomass-derived tars;black liquor tars.
 14. The process of claim 8, further comprisingfeeding one or more additional hydrocarbon feedstocks to the third-stagehydroprocessing reactor of contacting step (d), the one or moreadditional hydrocarbon feedstocks comprising hydrocarbonaceous materialsderived from thermal tars; bitumen; coke oven tars; asphaltenics; coalgasification tars; biomass-derived tars; black liquor tars.
 15. Anintegrated process for upgrading crude shale-derived oils produced byoil shale retorting or by in situ extraction or combinations thereofthat includes the following steps: a. contacting the whole shale oil andhydrogen in a first-stage hydroprocessing reactor containinghydrogenation catalysts to saturate diolefins contained in the wholeshale oil and recovering an effluent from the first-stagehydroprocessing reactor; b. feeding the effluent from the first-stagehydroprocessing reactor of step (a) without phase separation to asecond-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; c. fractionating the effluent from the second-stagehydroprocessing reactor of step (b) into a partially hydrotreatedfraction comprising naphtha, kerosene and diesel and a partiallyhydrotreated bottoms fraction comprising gas oil and residuum; d.feeding the partially hydrotreated fraction to a third-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andhydrodearomatization of the partially hydrotreated fraction andrecovering an effluent from the third-stage hydroprocessing reactor; e.feeding the partially hydrotreated bottoms fraction to a fourth-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the partially hydrotreated bottoms fraction andrecovering an effluent from the fourth-stage hydroprocessing reactor; f.processing the effluents from the third-stage hydroprocessing reactor ofstep (d) and the fourth-stage hydroprocessing reactor of step (e) in aseparation train to recover two or more hydrocarbon fractions.
 16. Theprocess of claim 15, wherein the two or more hydrocarbon fractionsinclude at least one of naphtha, kerosene, diesel, and a residuumfraction.
 17. The process of claim 15, wherein the catalyst in the firststage reactor comprises an alumina base extruded catalyst with nickeland molybdenum as active metals; the catalyst in the second stagereactor comprises an alumina base spheroidal catalyst with nickel andmolybdenum as active metals; the catalyst in the third stage reactorcomprises a layered catalyst comprised of a layer of amorphous basemetal Type II extruded catalyst with an organic compound and nickel andmolybdenum as active metals and a layer of base metal extruded catalystcontaining both amorphous and zeolitic components with nickel andtungsten as active metals; and the catalyst in the fourth stage reactorcomprises an amorphous base metal Type II extruded catalyst with anorganic compound and nickel and molybdenum as active metals.
 18. Theprocess of claim 15, further comprising: operating the first stagereactor at reaction conditions comprising a temperature in the rangefrom about 100° C. to about 250° C., a hydrogen partial pressure in therange from about 400 to about 500 psi, and a liquid hourly spacevelocity in the range from about 2 to about 6 L per hour per L catalyst;operating the second stage reactor at reaction conditions comprising atemperature in the range from about 200° C. to about 440° C., a hydrogenpartial pressure in the range from about 400 to about 2600 psi, and aliquid hourly space velocity in the range from about 0.5 to about 5.0 Lper hour per L catalyst; operating the third stage reactor at reactionconditions comprising a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst; and operating thefourth stage reactor at reaction conditions comprising a temperature inthe range from about 280° C. to about 440° C., a hydrogen partialpressure in the range from about 800 to about 2600 psi, and a liquidhourly space velocity in the range from about 0.3 to about 4.0 L perhour per L catalyst.
 19. An integrated process for upgrading crudeshale-derived oils produced by oil shale retorting or by in situextraction or combinations thereof that includes the following steps: a.contacting the whole shale oil and hydrogen in a first-stagehydroprocessing reactor containing catalysts to saturate diolefinscontained in the whole shale oil and recovering an effluent from thefirst-stage hydroprocessing reactor; b. feeding the effluent from thefirst-stage hydroprocessing reactor of step (a) without phase separationto a second-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; c. fractionating the effluent from the second-stagehydroprocessing reactor of step (b) into a partially hydrotreatedfraction comprising naphtha, kerosene and diesel and a partiallyhydrotreated bottoms fraction comprising gas oil and residuum; d.feeding the partially hydrotreated bottoms fraction to a third-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the partially hydrotreated bottoms fraction andrecovering an effluent from the third-stage hydroprocessing reactor; e.mixing the partially hydrotreated fraction and the effluent from thethird-stage hydroprocessing reactor to a mixture; f. feeding the mixtureto a fourth-stage hydroprocessing reactor having one or more beds ofcatalyst to perform hydrodenitrogenation, hydrodesulfurization,hydrodeoxygenation, and aromatics saturation of the mixture andrecovering an effluent from the fourth-stage hydroprocessing reactor; g.processing the effluent from the fourth-stage hydroprocessing reactor ina separation train to recover two or more hydrocarbon fractions.
 20. Theprocess of claim 19, wherein the two or more hydrocarbon fractionsinclude at least one of naphtha, kerosene, diesel, and a residuumfraction.
 21. The process of claim 19, wherein the catalyst in the firststage reactor comprises an alumina base extruded catalyst with nickeland molybdenum as active metals; the catalyst in the second stagereactor comprises an alumina base spheroidal catalyst with nickel andmolybdenum as active metals; the catalyst in the third stage reactorcomprises an amorphous base metal Type II extruded catalyst with anorganic compound and nickel and molybdenum as active metals; and thecatalyst in the fourth stage reactor comprises a layered catalystcomprised of a layer of amorphous base metal Type II extruded catalystwith an organic compound and nickel and molybdenum as active metals anda layer of base metal extruded catalyst containing both amorphous andzeolitic components with nickel and tungsten as active metal.
 22. Theprocess of claim 19, further comprising: operating the first stagereactor at reaction conditions comprising a temperature in the rangefrom about 100° C. to about 250° C., a hydrogen partial pressure in therange from about 400 to about 500 psi, and a liquid hourly spacevelocity in the range from about 2 to about 6 L per hour per L catalyst;operating the second stage reactor at reaction conditions comprising atemperature in the range from about 200° C. to about 440° C., a hydrogenpartial pressure in the range from about 400 to about 2600 psi, and aliquid hourly space velocity in the range from about 0.5 to about 5.0 Lper hour per L catalyst; operating the third stage reactor at reactionconditions comprising a temperature in the range from about 280° C. toabout 440° C., a hydrogen partial pressure in the range from about 800to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.3 to about 4.0 L per hour per L catalyst; and operating thefourth stage reactor at reaction conditions comprising a temperature inthe range from about 280° C. to about 440° C., a hydrogen partialpressure in the range from about 800 to about 2600 psi, and a liquidhourly space velocity in the range from about 0.3 to about 4.0 L perhour per L catalyst.
 23. An integrated process for upgrading crudeshale-derived oils produced by oil shale retorting or by in situextraction or combinations thereof that includes the following steps: a.contacting the whole shale oil and hydrogen in a first-stagehydroprocessing reactor containing catalysts to saturate diolefinscontained in the whole shale oil and recovering an effluent from thefirst-stage hydroprocessing reactor; b. feeding the effluent from thefirst-stage hydroprocessing reactor of step (a) without phase separationto a second-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; c. feeding the effluent from the second-stage hydroprocessingreactor of step (b) without phase separation to a third-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the effluent from the second-stagehydroprocessing reactor and recovering an effluent from the third-stagehydroprocessing reactor; d. fractionating the effluent from thethird-stage hydroprocessing reactor of step (c) into a partiallyhydrotreated fraction, comprising naphtha, kerosene and diesel, and apartially hydrotreated vacuum gas oil fraction; e. feeding the partiallyhydrotreated vacuum gas oil fraction to a fourth-stage hydroprocessingreactor having one or more beds of catalyst to perform hydrocracking ofthe partially hydrotreated vacuum gas oil fraction and recovering aneffluent from the fourth-stage hydroprocessing reactor; f. feeding theeffluent from the fourth-stage hydroprocessing reactor to thefractionation step (d); wherein: the catalyst in the first stage reactorcomprises an alumina base extruded catalyst with nickel and molybdenumas active metals; the catalyst in the second stage reactor comprises analumina base spheroidal catalyst with nickel and molybdenum as activemetals; the catalyst in the third stage reactor comprises a layer ofamorphous base metal Type II extruded catalyst with an organic compoundand nickel and molybdenum as active metals and a layer of base metalextruded catalyst containing both amorphous and zeolitic components withnickel and tungsten as active metal; and the catalyst in the fourthstage reactor comprises a base metal extruded catalyst containing bothamorphous and zeolitic components with nickel and tungsten as activemetal.
 24. The process of claim 23, wherein the two or more hydrocarbonfractions include at least one of naphtha, kerosene, diesel, and aresiduum fraction.
 25. The process of claim 23, further comprising:operating the first stage reactor at reaction conditions comprising atemperature in the range from about 100° C. to about 250° C., a hydrogenpartial pressure in the range from about 400 to about 500 psi, and aliquid hourly space velocity in the range from about 2 to about 6 L perhour per L catalyst; operating the second stage reactor at reactionconditions comprising a temperature in the range from about 200° C. toabout 440° C., a hydrogen partial pressure in the range from about 400to about 2600 psi, and a liquid hourly space velocity in the range fromabout 0.5 to about 5 L per hour per L catalyst; operating the thirdstage reactor at reaction conditions comprising a temperature in therange from about 280° C. to about 440° C., a hydrogen partial pressurein the range from about 800 to about 2600 psi, and a liquid hourly spacevelocity in the range from about 0.3 to about 4.0 L per hour per Lcatalyst; and operating the fourth stage reactor at reaction conditionscomprising a temperature in the range from about 330° C. to about 400°C., a hydrogen partial pressure in the range from about 1200 to about2600 psi, and a liquid hourly space velocity in the range from about 0.7to about 1.5 L per hour per L catalyst.
 26. An integrated process forupgrading crude shale-derived oils produced by oil shale retorting or byin situ extraction or combinations thereof that includes the followingsteps: a. contacting the whole shale oil and hydrogen in a first-stagehydroprocessing reactor containing catalysts to saturate diolefinscontained in the whole shale oil and recovering an effluent from thefirst-stage hydroprocessing reactor; b. feeding the effluent from thefirst-stage hydroprocessing reactor of step (a) without phase separationto a second-stage hydroprocessing reactor operated in an upflow mode andcontaining catalysts to perform hydrodemetallization and saturation ofmono-olefins in the effluent from the first-stage hydroprocessingreactor and recovering an effluent from the second-stage hydroprocessingreactor; c. feeding the effluent from the second-stage hydroprocessingreactor of step (b) without phase separation to a third-stagehydroprocessing reactor having one or more beds of catalyst to performhydrodenitrogenation, hydrodesulfurization, hydrodeoxygenation, andaromatics saturation of the effluent from the second-stagehydroprocessing reactor and recovering an effluent from the third-stagehydroprocessing reactor; d. fractionating the effluent from thethird-stage hydroprocessing reactor of step (c) into a partiallyhydrotreated fraction, comprising naphtha, kerosene and diesel, and apartially hydrotreated vacuum gas oil fraction; e. feeding the partiallyhydrotreated vacuum gas oil fraction to a fourth-stage hydroprocessingreactor having one or more beds of catalyst to perform hydrocracking ofthe partially hydrotreated vacuum gas oil fraction and recovering aneffluent from the fourth-stage hydroprocessing reactor; f. feeding theeffluent from the fourth-stage hydroprocessing reactor to thefractionation step (d); operating the first stage reactor at reactionconditions comprising a temperature in the range from about 100° C. toabout 250° C., a hydrogen partial pressure in the range from about 400to about 500 psi, and a liquid hourly space velocity in the range fromabout 2 to about 6 L per hour per L catalyst; operating the second stagereactor at reaction conditions comprising a temperature in the rangefrom about 200° C. to about 440° C., a hydrogen partial pressure in therange from about 400 to about 2600 psi, and a liquid hourly spacevelocity in the range from about 0.5 to about 5 L per hour per Lcatalyst; operating the third stage reactor at reaction conditionscomprising a temperature in the range from about 280° C. to about 440°C., a hydrogen partial pressure in the range from about 800 to about2600 psi, and a liquid hourly space velocity in the range from about 0.3to about 4.0 L per hour per L catalyst; and operating the fourth stagereactor at reaction conditions comprising a temperature in the rangefrom about 330° C. to about 400° C., a hydrogen partial pressure in therange from about 1200 to about 2600 psi, and a liquid hourly spacevelocity in the range from about 0.7 to about 1.5 L per hour per Lcatalyst.